U.S. patent number 7,964,699 [Application Number 12/311,664] was granted by the patent office on 2011-06-21 for polymer stream transfer.
This patent grant is currently assigned to Ineos Manufacturing Belgium NV. Invention is credited to Daniel Marissal, Brent R Walworth.
United States Patent |
7,964,699 |
Marissal , et al. |
June 21, 2011 |
Polymer stream transfer
Abstract
Process for heating a polymer-containing stream being
transferred from a polymerization reactor to a degassing vessel
operating at a pressure between 6 bara and 12 bara. The process
includes passing the stream through a heater having a transfer line
for the stream and a device for heating the transfer line. The
pressure drop in the heater is between 5% and 50% of the total
pressure drop between the polymerization reactor and the entry to
the degassing vessel. The pressure drop across the length of the
heater is less than 0.5 barh per tonne of polymer, and the average
Reynolds number across the cross-section of the stream at any point
along the length of the transfer line of the heater is greater than
500,000, such that at least 90 mol% of the hydrocarbon fluids
withdrawn from the polymerization reactor operation are vaporized
before entry into the degassing vessel.
Inventors: |
Marissal; Daniel (Carry le
Rouet, FR), Walworth; Brent R (Sint-Niklaas,
BE) |
Assignee: |
Ineos Manufacturing Belgium NV
(Antwerpen, BE)
|
Family
ID: |
38664458 |
Appl.
No.: |
12/311,664 |
Filed: |
October 4, 2007 |
PCT
Filed: |
October 04, 2007 |
PCT No.: |
PCT/EP2007/008602 |
371(c)(1),(2),(4) Date: |
April 08, 2009 |
PCT
Pub. No.: |
WO2008/043473 |
PCT
Pub. Date: |
April 17, 2008 |
Prior Publication Data
|
|
|
|
Document
Identifier |
Publication Date |
|
US 20100036078 A1 |
Feb 11, 2010 |
|
Foreign Application Priority Data
|
|
|
|
|
Oct 12, 2006 [EP] |
|
|
06255259 |
Oct 12, 2006 [EP] |
|
|
06255272 |
Oct 12, 2006 [EP] |
|
|
06255273 |
|
Current U.S.
Class: |
528/501;
528/502C; 422/144; 528/503; 422/145 |
Current CPC
Class: |
B01J
8/0015 (20130101); B01J 8/005 (20130101); B01J
19/1837 (20130101); B01J 19/2425 (20130101); B01J
19/2435 (20130101); B01J 19/2405 (20130101); B01J
2219/00247 (20130101); B01J 2208/0053 (20130101); B01J
2219/00159 (20130101); B01J 2208/00557 (20130101); B01J
2208/00548 (20130101); B01J 2219/00252 (20130101); B01J
2219/00094 (20130101); B01J 2219/00164 (20130101); B01J
2208/00539 (20130101); B01J 2219/00162 (20130101); B01J
2208/00663 (20130101); B01J 2208/00176 (20130101); B01J
2219/00038 (20130101); B01J 2208/00212 (20130101); B01J
2219/00166 (20130101) |
Current International
Class: |
C08F
6/00 (20060101); C08F 6/10 (20060101) |
Field of
Search: |
;528/501-503,502C
;159/2.1 ;422/144,145 |
References Cited
[Referenced By]
U.S. Patent Documents
Foreign Patent Documents
|
|
|
|
|
|
|
1415999 |
|
May 2004 |
|
EP |
|
2 028 836 |
|
Mar 1980 |
|
GB |
|
2 157 701 |
|
Oct 1985 |
|
GB |
|
WO 2006015807 |
|
Feb 2006 |
|
WO |
|
Other References
PCT International Preliminary Report on Patentability; PCT
International No. PCT/EP2007/008602; Int'l Filing Date Oct. 4,
2007. cited by other.
|
Primary Examiner: Cano; Milton I
Assistant Examiner: Fink; Brieann R
Attorney, Agent or Firm: Nixon & Vanderhye
Claims
The invention claimed is:
1. Process for heating a polymer-containing stream being
transferred from a polymerization reactor to a degassing vessel
operating at a pressure between 6 bara and 12 bara, comprising
passing the stream through a heater comprising a transfer line for
the stream and means for heating the transfer line, wherein the
pressure drop in the heater is between 5% and 50% of the total
pressure drop between the polymerization reactor and the entry to
the degassing vessel, the pressure drop across the length of the
heater is less than 0.5 barh per tonne of polymer, and the average
Reynolds number across the cross-section of the stream at any point
along the length of the transfer line of the heater is greater than
500000, such that at least 90 mol % of hydrocarbon fluids withdrawn
from the polymerization reactor operation are vaporized before
entry into the degassing vessel.
2. Process according to claim 1, wherein prior to entering the
heater, the polymer-containing stream passes through a pressure
control valve which introduces a pressure drop of between 45% and
90% of the total pressure drop between the polymerization reactor
and the entry to the degassing vessel.
3. Process according to claim 2, wherein the pressure drop across
the pressure control valve is between 10 and 30 barg.
4. Process according to claim 2, wherein the pressure control valve
is located between the polymerization reactor and the heater.
5. Process according to claim 2, wherein the ratio of the pressure
drop across the pressure control valve to the pressure drop across
the heater is between 0.1 and 6.
6. Process according to claim 1, wherein the pressure P.sub.i at
the heater inlet is 5-30 bara.
7. Process according to claim 1, wherein the pressure P.sub.o at
the heater outlet is 5-12 bara.
8. Process according to claim 1, wherein the pressure drop across
the transfer line per unit length is between 0.01 bar/m and 0.2
bar/m.
9. Process according to claim 1, wherein the temperature of the
polymer-containing stream at the heater exit is 5-20.degree. C.
above the dew point of the stream.
10. Process according to claim 1, wherein the temperature of the
internal surface of the transfer line along its length is
maintained at 10.degree. C. or more below the softening point of
the polymer, where the softening point of the polymer is defined as
the Vicat softening temperature according to ASTM D1525, ISO
306.
11. Process according to claim 1, wherein the ratio of the stream
velocity at the outlet of the heater to that at the inlet,
V.sub.o/V.sub.i, is at least 0.8.
12. Process according to claim 11, wherein the inlet velocity
V.sub.i is at least 2 m/s.
13. Process according to claim 11, wherein the outlet velocity
V.sub.o is less than 80 m/s.
14. Process according to claim 1, wherein the pressure drop across
the transfer line per unit length is between 0.0125 bar/m and 0.04
bar/m.
15. Process according to claim 1, wherein the solids content of the
polymer-containing stream when it enters the heater is between 35
wt % to 70.
16. Process according to claim 1, wherein the stream withdrawn from
the reactor is concentrated prior to passing through the
heater.
17. Process according to claim 1, wherein the polymer-containing
stream contains active polymer.
18. Process according to claim 1, wherein the polymer-containing
stream is continuously withdrawn from the polymerization
reactor.
19. Process according to claim 1, wherein the pressure drop in the
heater is between 10 and 35% of the total pressure drop between the
polymerization reactor and the entry to the degassing vessel.
20. Process according to claim 2, wherein the pressure control
valve introduces a pressure drop of between 60% and 80% of the
total pressure drop between the polymerization reactor and the
entry to the degassing vessel.
21. Process according to claim 3, wherein the pressure drop across
the pressure control valve is between 15 and 25 barg.
22. Process according to claim 5, wherein the ratio of the pressure
drop across the pressure control valve to the pressure drop across
the heater is between 0.1 and 2.
23. Process according to claim 6, wherein the pressure P.sub.i at
the heater inlet is 10-25 bara.
24. Process according to claim 7, wherein the pressure P.sub.o at
the heater outlet is 7-11 bara.
25. Process according to claim 9, wherein the temperature of the
polymer-containing stream at the heater exit is 10-15.degree. C.
above the dew point of the stream.
26. Process according to claim 10, wherein the temperature of the
internal surface of the transfer line along its length is
maintained at 20.degree. C. or more below the softening point.
27. Process according to claim 11, wherein the ratio of the stream
velocity at the outlet of the heater to that at the inlet,
V.sub.o/V.sub.i, is between 1.4 and 3.
28. Process according to claim 12, wherein the inlet velocity
V.sub.i is at least 5 m/s.
29. Process according to claim 12, wherein the inlet velocity
V.sub.i is at least 8 m/s.
30. Process according to claim 13, wherein the outlet velocity
V.sub.o is less than 70 m/s.
31. Process according to claim 1, wherein the polymer-containing
stream is heated in the heater such that at least 98 mol % of the
hydrocarbon fluids withdrawn from the polymerization reactor are
vaporized before entry into the degassing vessel.
32. Process according to claim 1, wherein the Reynolds number at
any point within the transfer line of the heater is between 1.8
million and 5 million.
33. Process according to claim 15, wherein the solids content of
the polymer-containing stream when it enters the heater is between
50 wt % and 65 wt %.
34. Process according to claim 16, wherein the stream withdrawn
from the reactor is concentrated prior to passing through the
heater using a hydrocyclone.
Description
This application is the U.S. National Phase of International
Application PCT/EP2007/008602, filed Oct. 4, 2007, which designated
the U.S. PCT/EP2007/008602 claims priority to European Application
Nos. 06255259.1, 06255272.4, and 06255273.2, filed Oct. 12, 2006,
respectively. The entire content of these applications are
incorporated herein by reference.
The present invention concerns apparatus for improving the
degassing of polymers, particularly olefin polymers.
Polymerisation of olefins in which an olefin monomer and optionally
olefin comonomer are polymerised, usually in the presence of a
catalyst and/or a diluent, is well known. The polymer is withdrawn
from the polymerisation reactor together with the reagents and
inert hydrocarbons. The reagents and hydrocarbons need to be
recovered for economic, safety and environmental reasons, and many
processes for achieving this are known in the art. These processes
generally involve depressurising and devolatilising the
polymer-containing stream after it has been withdrawn from the
polymerisation reactor. The volatisation requirement is greatest in
processes having polymer withdrawn from the reactor with high
absorbed or free liquid hydrocarbon contents. These are typically
processes producing polymers with a significant low density
component or amorphous phase where the absorption of hydrocarbon is
high, and/or processes where the polymer is produced in the
presence of liquid hydrocarbons (reactive or inert).
The maximum capacity of commercial scale plant has increased
steadily over the years, and as production rates have increased the
potential cost impact of unreliability in any part of the process
has also significantly increased, impacting not only the polymer
unit itself but also upstream and downstream units. At the same
time, growing operating experience has led to operation of
increasingly high solids concentrations (loadings) of polymer
withdrawn from reactors. The increase in solids concentrations in
slurry polymerisation units has typically been achieved with
increased circulation velocities achieved for example by higher
reactor power requirements as illustrated by EP 432555 and EP
891990. This increase in solids loadings is desirable as it
increases reactor residence time for a fixed reactor volume and
also reduces downstream diluent treatment and recycling
requirements. However, the transfer of the product at high solids
loadings is more problematic and careful design and operating
practices are required to avoid polymer fouling and blockage
problems that would not have been experienced at lower solids
loadings.
During, and as a result of, the process of depressurising and
devolatilising the polymer stream withdrawn from the polymerisation
reactor, the temperature of the polymer reduces. It is well known
that the process of devolatilising and desorbing the polymer is
significantly enhanced by maintaining the polymer at as high a
temperature as possible. Thus, in slurry loop processes the
transfer line between the polymerisation reactor and the
depressurisation (degassing) vessel for the polymer stream is
usually heated. As an example of a typical process, in WO 04/031245
and WO 05/044871 the take-off line from a loop polymerisation
reactor comprises a flashline containing the drawn-off slurry,
surrounded by a conduit which is provided with a heated fluid such
as low-pressure steam in order to provide indirect heating to the
slurry. However it is also well known that the stickiness and
susceptibility of polymer being transferred to agglomerate and/or
foul transfer lines and vessels generally increases with
temperature, and problems of fouling or agglomeration are becoming
more significant with the increasing solids loadings in the
transfer system now being utilised as mentioned above. As a result,
careful design of the transfer line from the polymerisation reactor
is required in order to achieve sufficient heating to aid
devolatilisation without risking fouling or agglomeration of the
solid polymer.
A further example of a process for transferring polymer slurry is
described in U.S. Pat. No. 4,126,743, in which a continuous process
is described for introducing the slurry into successive heating
zones, the second zone having a larger diameter than the first. It
is stated that the pressure in the reactor is preferably 10-30
kg/cm.sup.2 g, as is the pressure at the inlet of the first heating
zone, indicating that there is no pressure drop between the reactor
and the first heating zone. The pressure at the outlet of the first
heating zone is stated to be preferably 5-27 kg/cm.sup.2 g, and the
pressure at the outlet of the second heating zone is preferably
0/1-7 kg/cm.sup.2 g.
Devolatilisation of the polymer stream causes the liquid phase of
the stream to vaporise, resulting in an increase in volume in the
transfer line and a consequent increase in stream velocity. However
if the velocity becomes too high it may exceed the sonic velocity
(the speed of sound in the medium), leading to disruption of the
flow. On the other hand, if the initial velocity is too low there
is an increased risk of fouling or agglomeration of the solid
polymer as mentioned above.
A further consideration is that in large plants the transfer line
has to be very long in order to allow sufficient heating to take
place, and the length can be sufficiently great to impact the
spatial planning of the plant. This can create a variety of
problems such as footprint of the hardware in the plant, and
control of the conditions inside the line. Often it is necessary to
heat a significant proportion of the transfer line length to
satisfy the heat input requirements. Thus it will be appreciated
that ensuring that the polymer stream arrives at the degassing
vessel at the desired temperature and pressure and with the minimum
of fouling/agglomeration is a significant technical challenge.
The present invention seeks to optimise heating of the polymer
during its transfer from the reactor to the degassing vessel and at
the same time to maintain reliable product transfer, by particular
construction of the transfer line between the polymerisation
reactor and the degassing vessel for the polymer stream.
Accordingly in a first aspect the invention provides a process for
heating a polymer-containing stream being transferred from a
polymerisation reactor to a degassing vessel, comprising passing
the stream through a heater comprising a transfer line for the
stream and means for heating the transfer line, wherein the
pressure drop in the heater is between 5% and 50%, preferably
between 10 and 35%, of the total pressure drop between the
polymerisation reactor and the entry to the degassing vessel.
We have found that since the pressure at the outlet of the heater
is usually fixed by the downstream conditions, an excessively high
pressure drop across the heater implies a higher pressure at the
heater inlet, and hence a higher initial temperature of the
polymer-containing stream. The higher the initial temperature of
the polymer-containing stream, the smaller the temperature
difference between the stream and the heater itself, and hence the
less efficient the heating of the stream.
The polymer-containing stream may be continuously withdrawn from
the polymerisation reactor. The stream may or may not contain
active polymer.
It will be appreciated that the polymer-containing stream passes
through some form of pipeline at all times from the moment it
leaves the polymerisation reactor until it enters the degassing
vessel. For the purposes of this invention, the heater is
considered to comprise the portion of pipeline from the beginning
of the heated section (or first of the heated sections) until the
end of the heated section (or last of the heated sections). In this
context the term "heater" as used hereinafter includes within its
scope the possibility of a number of heaters connected in series.
The outlet of the heater (or transfer line of the heater) is
considered to be at the end of the heated section of the line, and
the inlet of the heater is considered to be at the start of the
heated section of the line, where the heated section of the line
incorporates the single heater or multiple heaters in series. By
"line" is meant any form of conduit suitable for transporting a
polymer-containing stream containing solids, liquid and gas.
The pressure drop across the transfer line per unit length is
preferably between 0.01 bar/m and 0.2 bar/m, preferably between
0.0125 bar/m and 0.1 bar/m, most preferably between 0.0125 bar/m
and 0.04 bar/m. Excessive pressure drops across the transfer line
are not desirable, as they may lead to choked flow conditions.
Typical pressure at the heater inlet, P.sub.i, is 5-30 bara,
preferably 10-25 bara. A higher Pi is undesirable as it implies a
small pressure difference between the heater inlet and the reactor
outlet, thereby ensuring that the temperature of the
polymer-containing stream at the heater inlet is still high. This
reduces the temperature difference between the polymer-containing
stream and the heater, thereby reducing heat transfer.
Pressure at the outlet P.sub.o is typically 1.5-12 bara, preferably
7-11 bara.
It is preferred that between exiting the polymerisation reactor and
entering the degassing vessel, the polymer-containing stream passes
through a pressure control valve which introduces a pressure drop
which is typically between 45% and 90%, preferably 60% and 80%, of
the total pressure drop between the polymerisation reactor and the
entry to the degassing vessel. A typical pressure drop across the
pressure control valve is between 10 and 30 barg, more usually
between 15 and 25 barg. Preferably the pressure control valve is
located between the reactor and the inlet of the heater, in which
case the pressure drop across the pressure control valve is the
difference between the reactor outlet pressure and P.sub.i.
Typically the ratio of the pressure drop across the pressure
control valve to the pressure drop across the heater is between 0.1
and 6, preferably between 0.1 and 2, and more preferably between
0.2 and 2, for example between 0.2 and 0.5.
In the process of the invention it is preferred that the
temperature of the polymer-containing stream at the heater exit is
maintained above the dew point of the non-polymer part of the
stream, preferably 5-80.degree. C., most preferably 10-30.degree.
C., above the dew point. The temperature of the internal surface of
the transfer line (process-side wall temperature) at any point
along its length is preferably maintained at 5.degree. C. or more
below the softening point of the polymer, more preferably
10.degree. C. or more below the softening point. The softening
point of the polymer is defined as the Vicat softening temperature
according to ASTM D1525, ISO 306. In the case where the heater is
positioned between a polymerisation reactor and a degassing vessel,
the temperature of the internal surface of the transfer line may be
maintained above the temperature of the reactor. For a polymer
having a density of 935-945 kg/m.sup.3 the process-side wall
temperature is typically controlled between 75 and 130.degree. C.,
preferably between 85 and 105.degree. C. For a polymer having a
density of 955-965 kg/M.sup.3 the process-side wall temperature is
typically controlled between 80 and 135.degree. C., preferably
between 95 and 110.degree. C.
The Vicat softening temperature according to ASTM D1525, ISO 306 is
the temperature at which a flat-ended needle penetrates a specimen
of the polymer to a depth of 1 mm under a load of ION. The
temperature reflects the point of softening to be expected when a
material is used in an elevated temperature application. A test
specimen, which is between 3 mm and 6.5 mm thick and at least 10 mm
in width and length, is placed in the testing apparatus (for
example a ROSAND ASP 6 HDT/VICAT System) so that the penetrating
needle, which has a cross-sectional area at its tip of 1 mm.sup.2,
rests on the specimen surface at least 1 mm from the edge. A load
of 10 N is applied to the specimen. The specimen is then lowered
into an oil bath at 23.degree. C. The bath is raised at a rate of
50.degree. C. per hour until the needle penetrates 1 mm; the
temperature at which this occurs is the Vicat softening
temperature.
The outlet temperature of the transfer line--specifically, the
internal wall temperature of the line at its outlet immediately
downstream of the final heated portion--is preferably maintained a
temperature higher than the temperature of the internal wall of the
line at its inlet immediately upstream of the first heated portion,
more preferably at least 5.degree. C. higher than the inlet
temperature.
For the process of the invention it is very desirable to maintain
the average velocity of the polymer-containing stream sufficiently
high to avoid the risk of fouling or blockages. By "average" is
meant the average velocity across the entire cross-section of the
stream at any one point along the length of the transfer line.
Hence it is preferred that the average velocity at the inlet
V.sub.i is at least 2 m/s, preferably at least 5 m/s and more
preferably at least 8 m/s. Typical values for V.sub.i are 3-20 m/s.
It is also desirable to maintain the velocity below the sonic
velocity. Accordingly it is preferred that the average velocity at
the outlet V.sub.o is less than 80 m/s, preferably less than 70
m/s. Preferably V.sub.o is at least 20 m/s; typical values for
V.sub.o are 30-80 m/s. It is preferred that the ratio of average
outlet velocity to average inlet velocity (where outlet and inlet
of the transfer line are as defined above), V.sub.o/V.sub.i, is at
least 1.1, typically between 1.2 and 15, preferably between 1.4 and
10, most preferably between 1.5 and 4.
In a preferred embodiment the process of the invention causes a
polymer-containing stream to be heated such that at least 90 mol %,
preferably at least 98 mol % and most optimally 100 mol % of the
hydrocarbon fluids withdrawn from the polymerisation reactor
operation are vapourised before entry into the degassing vessel.
The degassing vessel preferably operates at a pressure greater than
2 bara, most preferably between 6 bara and 12 bara whilst
maintaining a pressure drop across the length of the heater of less
than 0.5 barh per tonne of polymer, most preferably between 0.1
barh/te and 0.3 barh/te. It has been found that this optimised low
pressure drop per unit of production can be reliably operated even
at high solids loadings at entry into the heater. It is preferred
that the solids content of the polymer-containing stream is between
35 wt % to 70 wt %, most preferably between 50 wt % and 65 wt %,
when the stream enters the heater, and it is also preferred that
the velocity of the stream at the entry into the heater does not
vary by more than 15%, preferably no more than 5%, in any 30 second
period. One way in which this can be achieved is by using a
continuous rather than discontinuous off-take from the
polymerisation reactor. Such high solids loading operation combined
with the expanding diameter heater enables the pressure drop of the
heater to be minimised.
The average Reynolds number across the cross-section of the stream
at any point along the length of the transfer line of the heater
should always be greater than 500000, preferably between 1 million
and 10 million, most preferably between 1.8 million and 5
million.
The heating and depressurisation of the polymer stream as it
travels along the transfer line to the degassing vessel causes a
progressive vaporisation of the liquid in the stream and a
resultant increase in velocity along the line. There are
conflicting requirements to satisfy when designing the transfer
line in order to ensure effective and reliable polymer transfer and
heat transfer. Whilst high velocities enhance heat transfer and
generally minimise fouling, they also lead to high pressure drops
along the line. It is important therefore to be able to minimise
transfer line length and the required heat transfer area whilst
achieving a sufficiently de-volatised polymer at an acceptable
temperature.
Regarding the construction of the heater itself, it is preferred
that the ratio of the transfer line's outlet diameter D.sub.o to
its inlet diameter D.sub.i, D.sub.o/D.sub.i, is greater than 1,
preferably between 1.2 and 10. Typically it is at least 1.3, and
commonly at least 1.4. However this ratio is preferably no more
than 4, and more preferably no more than 2, with a maximum of 1.9
being most preferred. We have found that increasing the diameter of
the transfer line along its length allows a greater range of
polymer-containing stream flowrates to be accommodated by the
heater. A relatively small diameter at the inlet permits a
relatively high velocity even at low flowrates, reducing the risk
of fouling; whilst a relatively larger diameter at the outlet can
avoid the risk of the velocity exceeding the sonic velocity even at
high flowrates. Having such a range of capacities is particularly
valuable during start-up and shut-down operations. In order to
reduce the risk of downstream blockages it is also preferred that
the outlet diameter D.sub.o. of the transfer line is smaller than
the solids outlet of the degassing vessel. D.sub.o is defined as
the internal diameter of the transfer line at its outlet, and
D.sub.i is the internal diameter of the transfer line at its inlet,
where the outlet and inlet of the transfer line are defined as
previously described.
The internal diameter D of the transfer line is preferably at least
20 mm, and more usually between 40 mm and 200 mm. Most preferred
are internal diameters between 60 mm and 150 mm.
The length L of the heater, and therefore the transfer line, is
preferably at least 20 m, more preferably at least 30 m, but
usually no more than 600 m. A preferred range of lengths is from 50
m to 500 m, more preferably from 70 m to 300 m.
It is preferred that the ratio of the length L of the transfer line
to its average internal diameter D.sub.ave, L/D.sub.ave, is from
500 to 10000, preferably from 1500 to 3500, and more preferably
from 2000 to 3000. If the transfer line is constructed from a
number of sections each of different diameters, D.sub.ave is the
average internal diameter of those sections weighted according to
the length of each section; alternatively it may be calculated by
reference to the total internal volume V of the line, where
V=(.pi.D.sub.ave.sup.2.L)/4.
If the transfer line increases in diameter along its length, It is
preferred that the increase occurs in discrete steps rather than
continuously. Typically there are one, two or three increases in
diameter along the length of the pipe.
It is preferred that one or all of the sections of the line are
generally upright rather than mounted horizontally, so that the
line has a smaller footprint in the plant: in such a configuration,
the first section of the line preferably has its inlet at the
bottom such that the initial flow of material through the transfer
line is upwards. It is preferred that less than 20%, most
preferably less than 10% of the length of the transfer line is
horizontal, and optimally the line is constructed substantially
without horizontal sections. In one embodiment at least the inlet
and exit of the heated transfer line are oriented vertically such
that the inlet flow through the line is upwards and the outlet flow
from the line is downwards. In one embodiment of the invention, the
transfer line comprises a series of sections connected by bends
(elbows), which are typically U-shaped such that the line doubles
back on itself one or more times. The advantage of this
configuration is that it makes the transfer line more compact in
the plant. The sections between the elbows are usually straight.
The bends may be heated like the rest of the line, but usually--in
order to simplify the construction of the heater--they are not
heated. It is also generally preferred that any expansion in
diameter of the line occurs in an unheated section of the line;
therefore sections of the line may be of different diameters with
the increases in diameter occuring at one or more of the elbows,
preferably at the exit of the elbow such that the velocity reduces
at the exit of the elbow rather than at its inlet, and most
preferably at the exit of an elbow at the top of a vertically
heated section. The design of the expansion sections and the bends
in the transfer line is key to reliable operation without fouling.
The number of vertical or horizontal sections between elbows making
up the total transfer line may be from 2 up to 10, although 3 to 7
sections is more common.
The elbows of the transfer line may have varying degrees of
curvature. The radius of the curve defined by the elbow may be
expressed as a multiple of the diameter D of the line at that
point. The elbows typically have radii of between 3 D and 30 D,
with 5 D-20 D being most preferred to ensure reliable operation
without fouling whilst also minimising the footprint of the line.
As previously stated the elbows are preferably U-shaped, although
alternative options such as L-shaped elbows that permit a smooth
flow path are not excluded. Obviously a transfer line formed in
sections may employ a mixture of the above types of elbow, or
indeed elbows of other angles such as 60.degree. or
120.degree..
It has been found that the length of any one expansion section of
the transfer line should be greater than 0.25 D, preferably between
0.5 D and 10 D, most preferably between 0.75 D and 3 D. It is
preferred that each expansion section is located immediately
upstream or downstream of an elbow, preferably immediately
downstream of an elbow. It is also preferred that the expansion is
concentric, although other expansion geometries are also
possible.
The total specific heat transfer area of the transfer line, which
is its outer surface area in contact with the heating means, is
preferably at least 0.5 m.sup.2 of heat transfer area per tonne/h
of production of polymer, typically between 0.7 and 10, more
preferably between 1 and 5, most preferably between 1.5 and 3.5 of
heat transfer area per tonne/h of production of polymer.
It is preferred that the heater inlet is at approximately the same
elevation as the exit of the polymerisation reactor to which it is
linked, preferably the transfer line from the polymerisation
reactor to the heater inlet is essentially horizontal.
It is most preferred that the exit of the transfer line (at the
degassing vessel entry point) is at a higher elevation than the
inlet of the transfer line and/or the outlet of the polymerisation
reactor.
The means for heating the transfer line usually comprises a jacket
surrounding the line. The heater jacket may be in the form of an
electrical heater, but it is preferred that it is in the form of a
concentric pipe surrounding the line through which a heating fluid
is passed. The most commonly used heating fluid is steam. It has
been found that conditions can be best optimised by using
desuperheated steam as the heating medium, particularly where the
maximum saturation temperature is 0-30.degree. C., and preferably
no greater than 10.degree. C., below the softening point of the
polymer being heated. Whichever form it takes, the jacket may
either provide the same heat input along the entire length of the
transfer line, or may provide differential heating at different
parts of the line. It is also possible that portions of the line
(such as bends) are unheated, as discussed above. We have found
that the optimum heat input along the length of the transfer line
is achieved with a design such that temperature of the heating
medium (or the internal wall temperature of the line) is greater at
the inlet to the line than at its exit. Accordingly, as the vapour
fraction in the polymer-containing stream as it passes along the
line increases, it is preferred that the heating medium temperature
(or the internal wall temperature of the line) is decreased. This
can be done in a continuously graded manner, or in a number of
discrete steps by means of sections of different temperature. Most
preferred however is a jacket which operates at different
temperatures in different parts of the line, usually by having
independent supplies of the heating medium for each section where a
different temperature is required.
In a preferred embodiment of the invention, the transfer line is
heated by a concentric pipe using steam as a heating medium. The
outlet temperature of the transfer line is preferably controlled
using the steam flowrate: for a given steam temperature this has
the benefit of enabling control of the transfer line wall
temperature, to ensure a lower temperature at low polymer stream
flowrates and a higher temperature at higher flowrates when
velocities are higher.
One way of increasing further the temperature of the
polymer-containing stream itself at the transfer line outlet (apart
from increasing the energy input into the heater) is to increase
the solids content of the stream. This can be done by increasing
the solids content of the stream withdrawn from the polymerisation
reactor and/or by using a solids concentrating device upstream of
the transfer line. The solids can carry more heat than the liquid
or gaseous components of the stream, thereby requiring a lower
input from the transfer line heater in order to achieve the desired
temperature.
The use, upstream of the transfer line, of a solids concentrating
device with an upstream diluent flush (as described in our patent
EP1118624) is a preferred embodiment of the present invention, and
enables the monomer concentration in the transfer line to be
minimised, thereby reducing the risk of fouling.
It is preferred that the pipe is easily separable along the length
of the heater to facilitate cleaning. Preferably the pipe is
flanged at 5-15 m intervals. In the case where heating is effected
using a jacket containing a heating fluid, it is preferred that the
heating fluid does not cover any flange.
In order to maximise heat transfer into the polymer-containing
stream, the pipe is preferably made from a material having a
thermal conductivity greater than 30 Wm.sup.-2K.sup.-1, preferably
greater than 40 Wm.sup.-2K.sup.-1. The pipe is typically seamless,
although seam welded pipe is preferred where high heat transfer is
required.
It is preferred that all of the polymer-containing stream exiting
the polymerisation reactor is passed through a single transfer line
according to the invention. Such a transfer line may be fed by one
or more withdrawal lines from the reactor. The stream withdrawn
from the reactor may be concentrated, preferably by gravitational
or centrifugal means, most preferably using a hydrocyclone, prior
to passing through the transfer line. It is also within the scope
of this invention to provide multiple parallel transfer lines to
accept the polymer-containing stream, each of which is arranged
according to the invention. In such a construction, not all of the
transfer lines may be required in service at any one time. In a
further embodiment the polymerisation reactor has a plurality of
withdrawal lines, each of which has its own transfer line. The
invention also includes within its scope the use of single or
parallel solids concentrating devices, with the usual arrangement
being one solids concentrating devices located upstream of each
transfer line.
In the parallel heater embodiment, it is preferred that when at
least two heaters are operating, the average stream velocity across
any cross-section of each heater's transfer line is maintained
between 2 and 100 m/s, most preferably between 10 and 70 m/s. The
performance of each transfer line can be monitored using parameters
including the steam flow into the heating jacket to measure the
heat input (duty) to the stream, the pressure difference across the
heater and the reactor pressure valve output to measure the flow or
flow ratio into each transfer line, the relationship between steam
flow and outlet temperature for each heater, mass balance of the
reactor to calculate the total flow into all the heaters, and the
difference between the vapour temperature at the heater outlet and
the dew point of the process stream. Pressure drops across the
transfer lines of each heater are preferably essentially the same
as in the single heater embodiment as described above.
In both single heater and multiple parallel heater embodiments, the
flowrate of the polymer-containing stream withdrawn from the
polymerisation reactor is preferably controlled using a pressure or
flow control valve, most preferably located between a solids
concentrating device and the transfer line heater inlet. The
control valve is designed to have a pressure drop of between 45%
and 90%, most preferably 50% and 80%, of the pressure drop between
the reactor and the entry to the first downstream vessel. The
heated transfer line is preferably designed to have a pressure drop
between 5% and 75%, most preferably between 10 and 35%, of the
pressure drop between the reactor and the entry to the degassing
vessel. The ratio of the pressure drop across the control valve to
the pressure drop across the heater is between 0.8 and 5, most
preferably between 1 and 2.
The polymer-containing stream may contain a vapour component as
well as a liquid component. Typically the vapour fraction of the
fluid component of the polymer-containing stream at the inlet of
the heater varies from 5 to 60 mol %. In one preferred embodiment
of the present invention where there is a pressure or flow control
valve upstream of the heater, and the vapour fraction of the stream
at the heater inlet is between 25 and 60 mol %. The vapour fraction
of the fluid component of the stream at the outlet of the heater
can vary from 70 to 100 mol %, it is typically 95-100 mol %, most
preferably greater than 99 mol %.
This invention can be applicable to any polymerization process (eg
gas-phase, slurry, or solution) containing a polymer stream that
needs to be heated to volatilize liquid during
depressurisation.
Processes for the co-polymerisation of olefins in the slurry phase
are well known in the art. Such processes can be conducted for
example by introducing the monomer and comonomer into a stirred
tank or continuous loop reactor comprising polyolefin and a
catalyst for the polymerisation. The reactor is typically
controlled to achieve a desired melt index and density for the
polymer at an optimum production and temperature.
Polyethylene slurry polyrnerisation processes typically withdraw
polymer from the polymerisation reactor with significant quantities
of liquid hydrocarbons, and the present invention is therefore
particularly relevant to such processes. The slurry in such
reactors typically comprises the particulate polymer, the
hydrocarbon diluent(s), (co)monomer(s), catalyst, chain terminators
such as hydrogen and other reactor additives. In particular the
slurry will comprise 20-75, preferably 30-70 weight percent based
on the total weight of the slurry of particulate polymer and 80-25,
preferably 70-30 weight percent based on the total weight of the
slurry of suspending medium, where the suspending medium is the sum
of all the fluid components in the reactor and comprises the
diluent, olefin monomer and any additives; the diluent can be an
inert diluent or it can be a reactive diluent such as a liquid
olefin monomer. Where the principal diluent is an inert diluent the
olefin monomer typically comprises 2-20 wt %, more particularly
4-10 wt % of the slurry.
Polymerisation is typically carried out at temperatures in the
range 50-125.degree. C. and at pressures in the range 1-100 bara.
The catalyst used can be any catalyst typically used for olefin
polymerisation such as chromium oxide, Ziegler-Natta or
metallocene-type catalysts. The product slurry comprising polymer
and diluent, and in most cases catalyst, olefin monomer and
comonomer can be discharged intermittently or continuously,
optionally using concentrating devices such as hydrocyclones or
settling legs to minimise the quantity of fluids withdrawn with the
polymer.
This invention is particularly related to polymerisation in a loop
reactor where the slurry is circulated in the reactor typically by
means of a pump or agitator. Liquid full loop reactors are
particularly well known in the art and are described for example in
U.S. Pat. Nos. 3,152,872, 3,242,150 and 4,613,484. The loop reactor
is of a continuous tubular construction comprising at least two,
for example four, vertical sections and at least two, for example
four horizontal sections. The heat of polymerisation is typically
removed using indirect exchange with a cooling medium, preferably
water, in jackets surrounding at least part of the tubular loop
reactor. The volume of the loop reactor can vary but is typically
in the range 20 to 170 m.sup.3.
In commercial plants, the particulate polymer is separated from the
diluent in a manner such that the diluent is not exposed to
contamination so as to permit recycle of the diluent to the
polymerization zone with minimal if any purification. Separating
the particulate polymer produced by the process of the present
invention from the diluent typically can be by any method known in
the art for example it can involve either (i) the use of
discontinuous vertical settling legs such that the flow of slurry
across the opening thereof provides a zone where the polymer
particles can settle to some extent from the diluent or (ii)
continuous product withdrawal via a single or multiple withdrawal
ports, the location of which can be anywhere on the loop reactor
but is usually adjacent to the downstream end of a horizontal
section of the loop. As previously discussed, the operation of
large diameter reactors with high solids concentrations in the
slurry minimises the quantity of the principal diluent withdrawn
from the polymerisation loop. Use of concentrating devices on the
withdrawn polymer slurry such as hydrocylones (single or in the
case of multiple hydrocyclones in parallel or series) further
enhances the recovery of diluent in an energy efficient manner
since significant pressure reduction and vaporisation of recovered
diluent is avoided.
The withdrawn, and preferably concentrated, polymer slurry is
usually depressurised prior to being transferred via the heater of
the present invention to a primary flash vessel.
The diluent and any monomer vapours recovered in the primary flash
vessel are typically condensed, preferably without recompression,
and reused in the polymerization process. The pressure of the
primary flash vessel is usually controlled to enable condensation
with a readily available cooling medium (e.g. cooling water) of
essentially all of the flash vapour prior to any recompression. The
pressure in said primary flash vessel is generally in the range
2-25 bara, more typically 5-20 bara and most often 6-11 bara. The
solid material recovered from the primary flash vessel is usually
passed to a secondary flash vessel to remove residual volatiles.
Alternatively the slurry may be passed to a flash vessel of lower
pressure than in the above mentioned primary vessel such that
recompression is needed to condense the recovered diluent. Use of a
high pressure flash vessel is preferred.
More specifically, an example of the type of polymerisation process
for which the invention is particularly useful is the continuous
polymerization of olefins, preferably alpha mono olefins, in an
reaction zone, preferably an elongated tubular closed loop. The
olefin(s) is continuously added to, and contacted with, a catalyst
in a hydrocarbon diluent. The monomer(s) polymerise to form a
slurry of solid particulate polymer suspended in the polymerisation
medium or diluent. The rate of withdrawal of polymer product is
controlled by a valve upstream of the heater of the invention.
The solids concentration in the slurry in the reactor will
typically be above 20 vol %, preferably about 30 volume %, for
example 20-40 volume %, preferably 25-35 volume % where volume % is
[(total volume of the slurry-volume of the suspending
medium)/(total volume of the slurry)].times.100. The solids
concentration measured as weight percentage, which is equivalent to
that measured as volume percentage, will vary according to the
polymer produced but more particularly according to the diluent
used. Where the polymer produced is polyethylene and the diluent is
an alkane, for example isobutane, it is preferred that the solids
concentration is above 30 wt % in particular above 40 wt %, for
example in the range 40-60 wt % preferably 45-55 wt % based on the
total weight of the slurry. We have found that for high solids
loadings, particularly above 40 wt %, reliable product withdrawal
and heating between the polymerisation reactor and the degassing
vessel (as evidenced by fouling, flow variations and/or heat
transfer) can be maintained within acceptable operating limits by
use of the heater of the invention.
This type of process may optionally be carried out in a multiple
reactor system. The second or any subsequent reactor of the
multiple reactor system may be another loop reactor or any reactor
for the polymerisation of olefins, for example a fluidised-bed
reactor. However, usually the second or any subsequent reactor of
the multiple reactor system is another loop reactor. Such multiple
reactor systems can be used to make monomodal or multimodal,
preferably multimodal polymers.
In the case of multiple reactors in series, a first reactor of the
series is supplied with catalyst or prepolymer and optionally the
cocatalyst in addition to the diluent and monomer, and each
subsequent reactor is supplied with, at least, monomer, in
particular ethylene and with the slurry arising from a preceding
reactor of the series, this mixture comprising the catalyst and a
mixture of the polymers produced in a preceding reactor of the
series. It is possible to supply a second reactor and/or, if
appropriate, at least one of the subsequent reactors with fresh
catalyst and/or cocatalyst. However, it is more common to introduce
the catalyst and the cocatalyst exclusively into a first
reactor.
In the case where the plant comprises at least two reactors in
series, the polymer of highest melt index and the polymer of lowest
melt index can be produced in two adjacent or non-adjacent reactors
in the series. Hydrogen is maintained at (i) a low (or zero)
concentration in the reactor(s) manufacturing the high molecular
weight components, e.g. hydrogen percentages including between
0-0.1 vol % and at (ii) a very high concentration in the reactor(s)
manufacturing the low molecular weight components e.g. hydrogen
percentages between 0.5-2.4 vol %. The reactors can equally be
operated to produce essentially the same polymer melt index in
successive reactors.
When such reactor systems produce polymers of molecular weights
less than 50 kDaltons or greater than 150 kdaltons there have in
the past been observed particular problems with reactor fouling and
agglomeration in the heater between the polymerisation reactor and
the degassing vessel. These problems can be accentuated by high
polymer solids concentrations in the heater. This is another
problem which can be ameliorated by use of the heater of the
present invention.
EXAMPLE 1
A polymerisation reactor polymerising ethylene and hexene comonomer
in isobutane diluent is operated at 41 bar pressure and at a
temperature of 95.degree. C. This discharges a polymer-containing
stream which is in the form of a slurry, the liquid component of
which essentially comprises 91 mol % isobutane, 8 mol % unreacted
ethylene and 1 mol % hexene. The solids content of the slurry is
about 40 wt %, comprising polyethylene having a density of 940
kg/m.sup.3 and also some unused catalyst. These are the same
conditions as in Example 1.
The slurry from the reactor passes through a pressure control valve
to lower the pressure, before entering the heater of the invention.
Conditions at the inlet to the heater are:
Temperature: 82.4.degree. C.
Pressure P.sub.i: 17.4 bara
Velocity V.sub.i: 10.7 ms.sup.-1
Reynolds number: 1.72 million Solids concentration of 40 wt %; of
the remaining fluid phase, 40 wt % is vapour and 60 wt %
liquid.
The heater is 187 m in length and comprises 3 straight and vertical
sections each of 58 m length, which are connected by 180.degree.
elbows; the elbows together contribute 13 m of the total length and
are not heated. All the sections have an internal diameter of 78 mm
and wall thickness of 5.5 mm, giving a value of L/D.sub.ave of
about 2397. The thermal conductivity of the entire pipe wall is
46.4 W/mK. The heater element is in the form of concentric outer
tubes extending along each of the 26 m sections, through which is
passed desuperheated steam.
The slurry passes through the heater at a rate of 15 tonne/hour.
The length and diameter of the pipe, the heat input in the heater
from the heating medium to the slurry, and the velocity and initial
solids content of the slurry are all calculated to ensure that the
heat transfer to the slurry during its passage through the heater
is sufficient to ensure that the liquid phase is entirely vaporised
by the time the slurry exits the heater. The slurry exits the
heater at a temperature of 76.degree. C., a pressure P.sub.o of 9
bara and a velocity V.sub.o of 63.3 m/s (V.sub.o/V.sub.i=5.9), with
a Reynolds number of 3.3 million. This equates to a pressure drop
across the heater of 0.045 bar/m. At this pressure, the dewpoint of
the vapour would be about 60.8.degree. C., so the stream is
15.degree. C. above the dewpoint, and the liquid is completely
vaporised with no danger of any condensation of the vapour as it
exits the heater.
The internal wall temperature of the heater is between 89.degree.
C. and 93.degree. C. throughout the length of the heater; this
compares with a softening point of the polymer of about 128.degree.
C. The heat transfer coefficient from steam to slurry across the
wall of the heater is calculated as 984 W/m.sup.2K.
EXAMPLE 2
In this Example, the heater has a lower L/D.sub.ave ratio, but also
an increasing diameter.
A polymerisation reactor polymerising ethylene and hexene comonomer
in isobutane diluent is operated at 40 bar pressure and at a
temperature of 95.degree. C. This discharges a polymer-containing
stream which is in the form of a slurry, the liquid component of
which essentially comprises 91 mol % isobutane, 8 mol % unreacted
ethylene and 1 mol % hexene. The solids content of the slurry is
about 40 wt %, comprising polyethylene having a density of 940
kg/m3 and also some unused catalyst.
The slurry from the reactor first passes through a hydrocyclone to
concentrate the solids to 50 wt % and then a pressure control valve
to lower the pressure, before entering the heater of the invention.
Conditions at the inlet to the heater are:
Temperature: 76.degree. C.
Pressure P.sub.i: 14.4 bara
Velocity V.sub.i: 16.6 ms.sup.-1
Reynolds number: 2 million
Solids concentration of 50 wt %; of the remaining fluid phase, 40
wt % is vapour and 60 wt % liquid.
The heater is 152 m in length and comprises 5 straight and vertical
sections each of 26 m length, which are connected by 180.degree.
elbows; the elbows together contribute 22 m of the total length.
Each of the first three sections has an internal diameter of 78 mm
whilst each of the remaining two sections has an internal diameter
of 102 mm. The single increase in diameter occurs at the exit of
the elbow connecting the third and fourth sections. Thus the heater
has a value of D.sub.o/D.sub.i of 1.33 and a value of L/D.sub.ave
of about 1730. The thickness of the 78 mm diameter pipe wall is 5.5
mm, and the thickness of the 102 mm diameter pipe wall is 6.0 mm.
As in Example 1 the thermal conductivity of the entire pipe wall is
46.4 W/mK, and the heater element is in the form of concentric
outer tubes extending along each of the 58 m straight sections,
through which is passed desuperheated steam.
The slurry passes through the heater at a rate of 20 tonne/hour.
The length and diameter of the pipe, the heat input from the
heater, and the velocity and initial solids content of the slurry
are all calculated to ensure that the heat transfer to the slurry
during its passage through the heater is sufficient to ensure that
the liquid phase is entirely vaporised by the time the slurry exits
the heater. The slurry exits the heater at a temperature of
80.degree. C., a pressure P.sub.o of 10 bara and a velocity V.sub.o
of 30 m/s (V.sub.o/V.sub.i=1.78), with a Reynolds number of 2.7
million. This equates to a pressure drop across the heater of 0.03
bar/m. At this pressure, the dewpoint of the vapour would be about
65.degree. C., so the stream is 15.degree. C. above the dewpoint,
and the liquid is completely vaporised with no danger of any
condensation of the vapour as it exits the heater.
The internal wall temperature of the heater is between 89.degree.
C. and 93.degree. C. throughout the length of the heater; this
compares with a softening point of the polymer of about 128.degree.
C. The heat transfer coefficient from steam to slurry across the
wall of the heater is calculated as 600 W/m.sup.2 K.
* * * * *